ED apparatus having a multiplicity of alternating anion selective and cation selective membranes was apparently first described by K. Meyer and W. Strauss in 1940 (Helv. Chim. Acta 23 (1940) 795-800). The membranes used in this early ED apparatus were poorly ion selective. The discovery of ion exchange (“IX”) membranes (e.g., in U.S. Pat. No. Re. 24,865) which had high ion permselectivity, low electrical resistance and excellent stability led rapidly to the invention of ED using such membranes (e.g., in U.S. Pat. No. 2,636,852) and to the growth of industries using such apparatus, for example, for desalting of brackish water, concentration of sea water, and deashing of cheese whey. During the last 40 years approximately 5000 ED plants have been installed on a world-wide basis.
The utility of ED continues to be limited, however, by several technical actors, particularly relatively low limiting current densities and deficiencies in removing poorly ionized substances. These limitations and deficiencies of prior art ED systems are discussed further below.
A. Limiting Current Density:
Because the IX membranes used in ED are highly selective to ions of one sign or the other, a substantial fraction of the ions passing through the membranes must reach the membrane wails by diffusion from the ambient solution through laminar flow layers which develop along the interfaces between the membranes and the solutions being depleted of ions (the “dilute or diluting solutions or streams” as they are known in the art). The maximum rate of diffusion of ions through the diluting solution occurs when the concentration of electrolyte at such membrane interfaces is essentially zero. The current density corresponding to such zero concentration at a membrane interface is referred to in the art as the limiting current density. To increase the limiting current density it is necessary to increase the rate of ion diffusion, for example, by reducing the thickness of the laminar flow layers by flowing the ambient solution rapidly by the membrane surfaces and/or by the use of turbulence promoters, and/or by increasing the temperature. Practical limiting current densities are genera in the range of 5,000 to 10,000 amperes per square meter for each kilogram-equivalent of salts per cubic meter of solution (that is, 0.5 to 1 amperes per square centimeter for each gram-equivalent of salts per liter). A typical brackish water has a concentration of salts of about 0.05 kg-eq/m.3 (that is about 0.05-eq/l or about 3000 parts per million (“ppm”)), and therefore has a limiting current density in the range of about 250 to 500 amperes per m.2 (0.025 to 0.05 amperes per cm2). In order to maximize the utilization of ED apparatus, it is desirable to operate at the highest possible current densities. However, as the limiting current density is approached, it is found that water is dissociated (i.e., “split”) into hydrogen ions and hydroxide ions at the interfaces between the (conventional) anion exchange (“AX”) membranes and the diluting streams. The hydrogen ions pass into the diluting streams while the hydroxide ions pass through the AX membranes and into the adjacent solutions which are being enriched in ions (the “concentrate, concentrated, concentrating or brine solutions or streams” as they are known in the art). Because brackish water may often contain calcium bicarbonate, there is also a tendency for calcium carbonate to precipitate at the surfaces of the (conventional) AX membranes which are in contact with the concentrating streams. This problem previously has been addressed by several techniques: by chemical or IX softening of the feed waters or the concentrating streams; by adding acid to the feed waters or the concentrating streams (with or without decarbonation); by nanofiltration (“NF”); or, by regularly reversing the direction of passage of the electric current thereby changing the concentrating streams to diluting streams (and the diluting streams to concentrating streams). See, e.g., U.S. Pat. No. 2,863,813. Of the above techniques, the most successful process has been the last mentioned process, namely reversing the electric current, which is referred to in the art as “electrodialysis reversal” (“EDR”).
The theory of limiting current in ED shows that in the case of sodium chloride solution, for example, the cation exchange (“CX”) membranes should reach their limiting current density at values which are about ⅔ rds that of the AX membranes. Careful measurements have shown that such is indeed the case. However, as the liming current density of (conventional) CX membranes is approached or exceeded, it is found that water is not split into hydroxide ions and hydrogen ions at the interfaces between such CX membranes and the diluting steams. The difference in behavior relative to the water splitting phenomenon of (conventional) AX and CX membranes at their respective limiting currents has been explained in recent years as catalysis of water splitting by weakly basic amines in the AX membranes. AX membranes which have only quaternary ammonium anion exchange groups (and no weakly basic groups) initially do not significantly split water as their limiting current is approached. Such behavior continues for only several hours, however, after which period water splitting begins and increases with time. It is found that the AX membranes then contain some weakly basic groups which have resulted from hydrolysis of quaternary ammonium groups. It is concluded that splitting of water at conventional AX membranes at or near their limiting current densities is an unfortunate phenomenon which is unavoidable for practical purposes.
The existence of limiting current in ED also means that in dilute solutions the liming current densities are relatively very low. For example, at a concentration of salts of about 0.005 kg-eq/m.3 (that is about 0.005 g-eq/l or about 300 ppm, a concentration typical of drinking water), the limiting current density is in the range of from about 25 to 50 amperes per m.2 (0.0025 to 0.005 amperes per m.2), i.e., the transfer of salts per unit area per unit time is very low (e.g., 50 to 100 grams of salt per hour per square meter). This problem seems first to have been addressed by W. Walters et al. in 1955 (Ind. Eng. Chem. 47 (1955) 61-67) by filling the diluting stream compartments in an ED stack (i.e., a series of AX and CX membranes) with a mixture of strong base and strong acid ion exchange (IX) granules. Since then many patents have issued on this subject, among them U.S. Pat. Nos. 3,149,061; 3,291,713; 4,632,745; 5,026,465; 5,066,375; 5,120,416; and 5,203,976, which patents are incorporated herein by reference. Two modes of operation using such filled-cell ED (known as EDI) have been identified. In the first mode, the IX granules serve as extensions of the membrane surface area thereby greatly increasing the limiting current density. In the second mode, a current density is applied which is very much greater than the limiting current density even with the presence of the IX granules. Under these circumstances, the rate of water splitting at membrane-diluting stream interfaces is very high and the IX granules are predominantly in the strong base and strong acid forms respectively. The apparatus in this mode is therefore best described as operating as continuously electrolytically regenerated (mixed bed) ion exchange. An intermediate mode may also be identified in which there is some water splitting but the IX granules are not predominantly in the strong base and strong acid forms respectively.
Most filled-cell ED (that is, EDI) systems operate in both modes, e.g., (1) in the same ED cell, the first mode near the entrance to the cell and the second mode near the et (2) in cells in flow series between a single pair of electrodes; or, (3) in separate stacks in flow series (each stack with its own pair of electrodes). Filled-cell ED is used to replace reverse osmosis or conventional, chemically regenerated IX systems, e.g., a strong acid CX column followed by a weakly basic AX column or, at least in part, a mixed bed IX column. In either of the latter cases, the CX and AX granules are chemically regenerated separately, e.g., with aqueous acidic solutions of sulfuric acid or hydrochloric acid and aqueous basic solutions of sodium hydroxide respectively. Precipitates of calcium carbonate, calcium sulfate and magnesium hydroxide are thereby not obtained. The columns of fine granules are effective filters for colloid matter which is rinsed off the granules during the chemical regeneration. In contrast, in the case of EDI, any calcium, bicarbonate and/or sulfate removed from the diluting stream occurs in a higher concentration in the concentrating stream, particularly when it is desired to achieve high recoveries of the diluting stream (which is the usual case). Such higher concentrations frequently result in precipitation in the concentrating stream. Furthermore, it is inconvenient (though technically possible) to back-wash the IX granules in a filled-cell ED apparatus thereby removing any colloidal matter which may have been filtered out.
These problems with EDI are generally solved by pretreatment processes, for example: (1) regenerable cation exchange for softening followed by regenerable anion exchange absorbents for colloid removal and/or bicarbonate removal; (2) ultrafiltration or microfiltration for colloid removal followed by EDR for softening and partial demineralization; or, (3) ultrafiltration or microfiltration for colloid removal followed by nanofiltration for softening or reverse osmosis for softening and partial demineralization.
As pointed out above, filled-cell ED is used to replace, at least in part, a mixed bed IX column. The latter, however, generally produces water having an electrical resistance of about 18 meg ohm-cm and silica concentrations near the present limits of detection. Such high performance by filled-cell ED (EDI) has been difficult to achieve until now.
B. Removal of Poorly Ionized Substances:
ED (including EDR) is used in many plants to deash cheese whey. Generally the natural whey is first concentrated to the range of 20 to 25 percent solids by weight. The current density (that is, the rate of removal of ash per unit area of membrane per unit time) during ED (or EDR) of such concentrated whey remains relatively high until about 50 to 60 percent of the ash is removed. The remaining ash behaves as if it is poorly ionized, perhaps associated or complexed with protein in the whey. An important market for deashed whey requires 90 percent or higher deashing. To deash from about a 40 percent ash level to a 10 percent ash level using ED (including EDR) may require much more apparatus contact time than to deash from 100 percent to 40 percent ash. This problem may be addressed by the more or less continuous addition of acid to the whey during deashing from 40 to 10 percent ash, the acid apparently freeing the ash from the protein. However, such added acid is rapidly removed by ED (including EDR), and the resulting high quantities of acid required to complete this process are therefore undesirable. The problem has also been addressed by removing about the first 60 percent of the whey ash using ED (including EDR) and removing most of the remaining 40 percent by ion exchange. The ion exchange apparatus for this application generally consists of a column of strong acid CX granules followed by a column of weak base AX granules. Considerable quantities of acid and base are required in this process to regenerate the IX granules.
As discussed above, electrodeionization (EDI) using filled ED cells is a very useful process for removing the last traces of ionic contaminants from water, but it could be significantly improved. These desired improvements include:    1) unproved product purity. The sources of impurities in EDI systems include but are not limited to:            (a) Back diffusion and electromigration of ionic contaminants through the ion exchange membranes driven by concentration differences and electric fields;        (b) Back diffusion of neutral weakly ionized species through polarized membranes; and,        (c) Electrodialysis of contaminant ions from membranes into the product water in the dilute stream manifolds.            2) Simplification of the equipment and controls needed to run traditional EDI. Today these EDI subsystems include brine recirculation, brine and electrolyte pH and conductivity control. Simplification should include lowering of equipment cost, reducing required operator expertise and shortening the time required for monitoring and adjusting equipment.    3) Reduction of the electric power consumption used for the EDI stack and pumps.    4) Improved resistance to scale formation in the concentrate streams.    5) Need for less pretreatment of the water before EDI.    6) Improved product water quality without sacrificing product water recovery.    7) Ability to operate intermittently and to produce excellent product immediately upon    8) Ability to operate with elevated solution delivery pressure thereby eliminating the need for a transfer pump.    9) Ability to reliably operate without external leaks and without external salt build up.
Most, if not all, of the above limitations and deficiencies of conventional EDI systems are either overcome or at least significantly improved upon by the improved apparatus and methods for electrodialysis according to the present invention. Other objects and advantages of the present invention will in part be obvious and will in part appear hereinafter. The invention accordingly comprises, but is not limited to, the apparatus and related methods, involving the several steps and the various components, and the relation and order of one or more such steps and components with respect to each of the others, as exemplified by the following description and the accompanying drawings. Various modifications of and variations on the apparatus and methods as herein described will be apparent to those skilled in the art, and all such modifications and variations are considered within the scope of the invention.